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Effect of SO2 on the amine-based CO2 capture solvent and improvement using ion exchange resins
International Journal of Greenhouse Gas Control 37 (2015) 38–45
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Effect of SO2 on the amine-based CO2 capture solvent and improvement using ion exchange resins Jubao Gao a,∗ , Shujuan Wang b,∗∗ , Jian Wang a,c , Lingdi Cao a,c , Shiwei Tang a , Yu Xia a a Beijing Key Laboratory of Ionic Liquids Clean Process, Key Laboratory of Green Process and Engineering, State Key Laboratory of Multiphase Complex Systems, Institute of Process Engineering, Chinese Academy of Sciences, Beijing 100190, China b Key Laboratory for Thermal Science and Power Engineering of Ministry of Education Department of Thermal Engineering, Tsinghua University, Beijing 100084, China c School of Chemistry and Chemical engineering, University of Chinese Academy of Sciences, Beijing 100049, China
a r t i c l e
i n f o
Article history: Received 28 July 2014 Received in revised form 12 January 2015 Accepted 4 March 2015 Keywords: Carbon dioxide Sulfur dioxide Pilot plant Heat stable salts Ion exchange resin Energy consumption
a b s t r a c t An amine-based CO2 capture solvent is developed to reduce energy requirement. Three campaigns, with 12 vol% CO2 and 18 vol% O2 , are conducted in a CO2 capture pilot plant to evaluate energy consumption and influence of SO2 on solvent characteristics, under the condition of prolonged operation. Monoethanolamine (MEA) is used in one campaign, and the other two campaigns employed new solvent, with and without 200 ppm SO2 in the simulated flue gas. The lost of free amine, including CO2 combines with amine (CCA), SO2 combines with amine (SCA) and amine degradation (AD), is analyzed. Experimental results show that the average energy consumption of MEA and novel absorbent is about 4.09 GJ/tCO2 and 2.96 GJ/tCO2 , respectively. This value climbs to 3.51 GJ/tCO2 for novel absorbent, with 200 ppm SO2 . Meanwhile, the pH is reduced, the viscosity increases and free amine is lost more, which reduces the overall CO2 reaction rate and the overall mass transfer coefficient. The presence of SO2 increases SCA and AD, and decreases CCA. The SCA becomes main factor after running 200 h. An industrial strong-alkaline anion exchange resin through several runs is used to decrease the SCA and improve the solvent characteristics. The heat stable salts (HSS), including formate, oxalate, glycolate, sulfite and sulfate, can be reduced to less than 1 wt% by using the ion exchange resin. The solvent viscosity after treatment is only one tenth of the original solvent at 313 K. The overall mass transfer coefficient KG after treatment is about 1.7 times that without treatment. © 2015 Elsevier Ltd. All rights reserved.
1. Introduction Amine scrubbing systems have been used to separate CO2 from natural gas in the chemical industry for several decades. The flue gas from coal-fired power plants, however, is quite different from that of natural gas, with large flow rates, high temperatures, low CO2 partial pressures, and the existence of SO2 , NOx and fly ash in the flow stream (Xu et al., 2014; Zhang et al., 2012). Therefore, removing CO2 from power plants results in significant energy consumption and reduction of electricity output (Bouillon et al., 2009; Knudsen et al., 2009; Razi et al., 2013). Many ways have been reported to minimize the CO2 capture energy requirement (Ahn et al., 2013; Fang et al., 2012; Gao et al., 2014; Warudkar et al., 2013a,b), solvent screening is another efficient method to improve
∗ Corresponding author. Tel.: +86 10 82544875; fax: +86 10 82544875. ∗∗ Corresponding author. Tel.: +86 10 62788668; fax: +86 10 62770209. E-mail addresses: [email protected] (J. Gao), [email protected] (S. Wang). http://dx.doi.org/10.1016/j.ijggc.2015.03.001 1750-5836/© 2015 Elsevier Ltd. All rights reserved.
the energy efficiency (Gao et al., 2011a; Kumar et al., 2014; Mathias et al., 2013; Pirngruber et al., 2013; Yu et al., 2011; Zhang et al., 2012). Recently, Li et al. (2013) found that about 380 patents have been published on CO2 capture by solvents. In this work, a novel blended amine solvent was developed and assessed in a pilot plant, with the long-term running. In addition, SO2 in the flue gas is one of the main factors causing free amine lost, which may lead to rise of CO2 capture energy consumption. In previous work, experiments have proved that SO2 in the flue gas results in serious amine degradation (AD), with many heat stable salts (HSS) produced (Gao et al., 2011a,b,c; Yang et al., 2013). Then, foaming and corrosion/erosion can be caused by the high levels of HSS (>1 wt% as amine) (Gao et al., 2012; Narendra and Anil, 2009). However, it is rarely to analysis the effect of SO2 on the CO2 combines with amine (CCA) and SO2 combines with amine (SCA), which leads to free amine lost. The upgraded FGD with more than 99% SO2 removal efficiency before CO2 capture was suggested by Rao and Ruben (2002) to solve the problems. But HSS and impurities are still accumulated in amine solutions. Hence, amine
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
Nomenclature Cam DCO2, L
amine concentration in the liquid phase (kmol/m3 ) CO2 diffusion coefficient in the amine solution
DCO2, w HCO2, w
(m2 /s) CO2 diffusion coefficient in water (m2 /s) Henry’s law coefficient of CO2 in the amine solution
Qstrip Qsens Qdes Qloss QT rov
(kPa m3 /kmol) second order rate constant of amine (m3 /kmol s) overall mass-transfer coefficient (kmol/(m2 h kPa)) gas-side mass-transfer coefficient (kmol/(m2 h kPa)) liquid-side mass-transfer coefficient (kmol/(m2 h kPa)) mass-transfer flux of the absorbed component CO2 (kmol/(m2 h)) average CO2 partial pressure (kPa) equilibrium CO2 partial pressure (kPa) CO2 partial pressure at the wetted wall column inlet (kPa) CO2 partial pressure at the wetted wall column outlet (kPa) latent heat of vaporization (kJ/mol CO2 ) sensible heat (kJ/mol CO2 ) heat of reaction (kJ/mol CO2 ) heat of loss (GJ/tCO2 ) total energy consumption (GJ/tCO2 ) overall CO2 reaction rate (kmol/m3 s)
Greeks L W ˛CO2
amine solution viscosity (mPa s) water viscosity (mPa s) CO2 loading (mol/mol)
k2 KG kg k g NCO2 PCO2, b P ∗ CO2 PCO2, in PCO2, out
reclamation is essential during CO2 capture to control the levels of HSS and remove impurities. Amine reclamation can be used to prevent potential secondary pollution with HSS removal and increases in the free amine concentration in a continuous absorption/regeneration process. Three prime technologies used commercially for solvent reclamation as thermal reclamation, electrodialysis, and ion-exchange. Thermal reclamation at the bottom of the stripper (Fig. 1) can result in higher reclamation costs and potential degradation of the amine due to the high temperatures (Xu and Rochelle, 2009). This is a useful method to remove non-volatile species. Electrodialysis is used to remove anions and cations, however, it can cause absorbent losses. Moreover, the ion-exchange membrane is relatively expensive especially in the presence of fly ash (Meng et al., 2008). Ion-exchange systems acts like electrodialysis disable to remove nonionized impurities. And ion-exchange resins must be used at low temperatures to reduce resin lost. Hence, ion-exchange systems all normally installed at the absorber inlet (Fig. 1). Ion-exchange has been widely used in coal-fired power plants to treat boiler feed-water, thus, it may be more easily accepted in coal-fired power plants for CO2 capture. The ion-exchange technology for removing HSS was patented by Keller (1991) and Chen et al. (2008). However, there is little information on the HSS type and removal efficiency. Meanwhile, these patents were only applied in removing H2 S and CO2 from natural gas. Although there is substantial information on novel solvents and methods of saving energy requirement, the information on the energy consumption analysis from pilot plant and effects of SO2 on absorbents, in particulate, improved methods are scant. Three
39
Table 1 Anion exchange resin properties.a Total exchange capacity (eq/L)
Moisture holding capacity (%)
Shipping weight (g/L)
Particle mean size (mm)
Uniformity coefficient
1.25
45–51
680
0.664
1.6
a
Data obtained from supplier except for particle size.
campaigns have been conducted in a pilot plant using simulated flue gas to assess CO2 capture energy consumption and effect of SO2 on CO2 capture solvent in this study. The main characteristics of amine aqueous solutions, such as viscosity, pH and bound amine, were analyzed after long-term operation, with and without 200 ppm SO2 . The AD, CCA and SCA of two campaigns were compared. An industrial grade anion exchange resin was selected to improve the characteristics of the amine-based CO2 absorbent. 2. Experimental 2.1. Materials The chemical reagents are industry grade, with concentrations of more than 98%. CO2 , SO2 , O2 and N2 , with more than 99.9% concentration, are used to simulate coal-fired flue gas. New solvent is a blend of a tertiary amine and additives. Aqueous amine solution samples are taken from the stripper bottom of the pilot plant every 100 h (Gao et al., 2011a,b,c). The character of four types of strong-alkaline and two types of weak-alkaline anion-exchange resin is compared, and the best is selected. Five solvent samples from the SO2 test are treated by the selected industrial grade strong base anion exchange resin from Rohm and Haas. The resin is made of pale yellow, translucent spherical beads with the main component being styrene divinylbenzene copolymer matrix with a dimethyl ethanol ammonium functional group. The anion is chloride as shipped and is converted to hydroxyl by percolation with an excess of 4 M NaOH solution and washed with deionized water. Table 1 shows the typical properties of the anion exchange resin. 2.2. Pilot plant operating conditions Three campaigns run in a CO2 capture pilot plant. Some important parameters are shown in Table 2. Two baseline campaigns operated, respectively, for 200 h and 500 h using 30 wt% MEA and novel solvent, with a feed gas of 12 vol% CO2 , 18 vol% O2 and N2 as the balance, and another campaign using the novel solvent, feed with an additional 200 ppm SO2 , operated for 430 h. The pilot plant is same as conventional amine-based CO2 recovery plant except for recycled flue gas from stripper top to absorber inlet (Gao et al., 2011a,c). The average flue gas flow rate of the two campaigns is about 86 Nm3 /h, and the average liquid solvent flow rate is around 0.6 m3 /h. The heat input to the stripper is supplied by electrical heater with the power of 60 kW. No fresh solvent and water are added into the system during the two campaigns. Thus, the CO2 removal efficiency is changed during the two campaigns because of free amine lost. The average temperature at the bottom of absorber and stripper is respectively around 46 ◦ C and 116 ◦ C. More details can be found from our previous work (Gao et al., 2011c). 2.3. Ion exchange procedure The resin is first percolated and held for 24 h with an excess of 4 M NaOH solution. Then it is washed three times and held for 3 h in deionized water. The pretreated resin is then put into ten 1 L beakers with 170 ml in each. 100 ml solvent samples are fed into
40
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
Fig. 1. A general flow diagram of CO2 capture with flue gas pretreatment and amine reclamation. Table 2 Experimental conditions. Items Unit
CO2 Vol%
O2 Vol%
SO2 ×10−6
Amine concentration mol/L
Flue gas flow rate Nm3 /h
Operating time h
Campaign 1 Campaign 2 Campaign 3
12.89 11.89 11.86
17.4 18.2 17.8
0 0 200
4.90 4.64 4.57
85.35 86.12 82.83
200 500 430
2.4. Analysis methods Dionex ICS-1000 and DX120 ion chromatographics are used to analyze the anions and cations. CO2 loading is determined by a total organic carbon analyzer (TOC-5000, Shimadzu). The solution viscosity is measured by a viscometer (SV-10, AND). The pH is measured on an automatic potentiometric titrator (809 Titrando, Metrohm, Switzerland). A wetted wall column with an interfacial area of about 41.45 cm2 is used to measure the kinetics of the 430 h samples before and after treatment. The column is described by Liu et al., (2009).
6
New Solvent MEA energy consumption ( GJ/t CO2)
each beaker. Five different samples from the SO2 tests (100 h, 200 h, 300 h, 400 h and 430 h samples) are held for 60 minutes in each beaker. The 430 h samples are fed into five beakers, each one lasting 10 min longer than the previous one (10, 20, 30, 40 and 50 min after reaction at room temperature). The mixture is stirred at a constant rate before the 10 ml samples is taken for analysis.
5
4
3
2 0
100
200
300
400
500
Time (h) Fig. 2. Energy consumption of different solvent.
3. Results and discussion 3.1. Energy consumption The total energy consumption, including energy loss of the whole CO2 capture system, the sensible heat, the latent heat of vaporization and the overall enthalpies of absorption, is determined experimentally by the following correlation: QT =
(tin − tout ) Coil oil Voil qCO2 ∗ 1000
(1)
where, tin and tout denote temperature of oil entering and leaving the reboiler, respectively. Coil , oil and Voil are the heat capacity, the density and the volume flow rate of oil with unit of kJ/(kg K), kg/m3 and m3 /h, respectively. qCO2 reprensents the mass flow rate of gas CO2 at the stripper top with unit of kg/h.
Results of the total energy consumption are shown in Fig. 2 and Fig. 3. It can be seen from Fig 2 that the total average energy consumption for 30 wt% MEA is about 4.09 GJ/tCO2 (3.53–4.88 GJ/tCO2 ), which is in good agreement with simulated results of Dave et al. (2009) and Cifre et al. (2009), but about 10% higher than that reported by Knudsen et al. (2009). For new solvent, however, it is about 2.96 GJ/tCO2 (2.46–3.15 GJ/tCO2 ), which is about 27% lower than that of MEA. After adding SO2 into the CO2 capture system (Fig. 3), this value has been increased to 3.51 GJ/tCO2 (1.69–5.05 GJ/tCO2 ). Before running of 150 h, campaign 3 has a comparable value to campaign 2. After running of 150 h, nevertheless, campaign 3 is gradually higher than campaign 2 due to the reduction of the CO2 removal efficiency (Gao et al., 2011c).
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
140
7.0 6.5
0ppm SO2
130
6.0
200ppm SO2
120
sensible heat heat of reaction heat of vaporization loss energy
110
5.5
100
5.0
Energy distribution %
energy consumption ( GJ/t CO2)
41
4.5 4.0 3.5 3.0 2.5
90 80 70 60 50 40
2.0
30
1.5
20 10
1.0 0
50
100
150
200
250
300
350
400
450
500
0
MEA
Time (h)
New Solvent
Absorption Solvent Fig. 3. Effect of SO2 on energy consumption. Fig. 4. Distribution of the total energy.
In addition, each component (loss heat, latent heat of vaporization, sensible heat and reaction enthalpy) can be expanded separately from the pilot plant experiments as following (Rochelle et al., 2002): QT = Qloss + Qsens + Qdes + Qstrip
(2)
Qdes = −Habs
(3)
Qsens
Qstrip =
(4)
2
Trich,str , ˛rich
n
Qloss = 3.6
(5)
18
Fresh solvent
14
8
0ppm SO2
6
[˛ti Ai (tsi − ta )] /qCO2 /1000
(6)
3.2. Influence of SO2 on the absorbent The HSS contents of the five SO2 samples are shown in Table 3. Most of the HSS is sulfates and sulfites because SO2 is absorbed conTable 3 HSS from CO2 capture pilot plant. Glycolate
Formate
Oxalate
105.7 170.2 271.2 396.8 375.4
22.1 78.4 157.9 248.4 335.5
Sulfate
Sulfite
mg/L 322.7 680.6 965.5 1302.2 1706.9
4 Treatment with 2 dry resins
Treatment with wet resins
0
where ˛ti and Ai are the total heat transfer coefficient and surface area of the i-th equipment with unit of w/(m2 K) and m2 , respectively. tsi and ta denote the surface average temperature of the i-th equipment and atmosphere temperature, respectively. The distributions of the four energy components for campaign 1 and 2 are illustrated in Fig. 4. It can be found that MEA has a comparable loss energy and sensible heat to new solvent. However, the heat of reaction and the latent heat of vaporization of new solvent are all lower than those of MEA, which implies that regenerating the new solvent requires low reboiler heat duty to break the chemical bounds between CO2 and amine and strip CO2 .
100 h 200 h 300 h 400 h 430 h
20
10
HHvap 2O
i=1
Sample
22
12
sat PH Trich,str xH2 O O ∗ PCO 2
200ppm SO2
24
16
CP = (˛rich −˛lean ) Cam
26
5007.5 12,144.0 19,454.7 25,723.2 33,837.0
5145.4 5118.2 5332.1 6095.5 8650.4
300
305
310
313K
315
320
325
330
335
340
Temperature (K) Fig. 5. Solvent viscosities.
tinuously. The effect of SO2 on the solvents is shown in the following sections. 3.2.1. Viscosity The viscosities of fresh amine, the 500 h sample and the 430 h sample are measured. As shown in Fig. 5, the solvent viscosities of the 500 h samples (0 ppm SO2 ) are similar to that of the fresh amine. The highest viscosity of the sample (200 ppm SO2 ), however, is observed. In general, amine concentration is proportional to the viscosity. Thus, two spent absorbents should cut down on the viscosity because of amine degradation, which indicates that the formation of HSS increases the viscosity especially in the presence of SO2 . Furthermore, the CO2 diffusion coefficient in solution DCO2, L can be calculated from the viscosity as (Prashanti et al., 2010; Versteeg and van Swaaij, 1988): DCO2, L = DCO2, W
0.8 W
(7)
L
DCO2, W = 2.35 × 10−6 exp(
−2119 ) T
(8)
42
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
[AM]. The former can be neglected because of rather low concentration. But the reduction of [OH− ] indicates the decrease of [AM], which results in the overall CO2 reaction rate decreases. The pH of the five samples before and after treatment is shown in Fig. 6. The pH increases significantly after the ion exchange resin treatment, which liberates part of bound amine and leads to an increase in the overall CO2 reaction rate as indicated by Eq. (12). At the same time, most of the impurity anions are removed by the resin so that the ionic strength of the solution decreases. As a result, Henry’s constant for the electrolyte solution is also reduced. All these results show an increase in the liquid-side mass transfer after the ion exchange treatment.
12.0
Before treatment for CO2 +200ppm SO2 CO2 only
11.5
After treatment for CO2 +200ppm SO2
11.0
pH
10.5
10.0
9.5
3.2.3. Free amine As an acidic gas, SO2 will combine with free amine, which will reduce the free amine concentration:
9.0 0
100
200
300
400
500
Time (h)
The Henry’s law constant of CO2 is obtained using (Gregory and Ralph, 1994):
I = (0.5)
Ci Zi2
(13) HCOO− ,
Fig. 6. pH at the bottom of the stripper.
log HCO2, L = h+ + h− + hg I + log (H∗ )
SO2 + H2 O + 2RNH2 → 2RNH3 + + SO3 2−
(9) (10)
2−
In effect, the organic anions, such as C2 O4 and C2 H3 O3 − , are also combined with free amine. Their influence is not considered in this study because of low concentration. The effect of H+ and OH− on the mass balance and charge balance are ignored because of rather low concentration. The forms of the absorbed CO2 are regarded as HCO3 − since the CO2 and CO3 2− are also small. The mass balance and charge balance equations can be approximated as follows: CR1 NH2 + CR
+ 1 NH3
= C1
(14)
where HCO2, L is Henry’s constant for the electrolyte solution and
CR2 NH2 + CR2 NH3 + = C2
(15)
H* is for the pure molecular solvent. h+ and h− are the van Krevelen coefficients for the cations and anions in the solution and is the coefficient for the dissolved gas. I is the ionic strength of the solution. Ci is the concentration of species i and Zi is the charge. According to Eqs. (7) and (8) DCO2, L decreases as the viscosity
CHCO− = (C1 + C2 )˛CO2
(16)
increases. And the accumulation of SO3 2− and SO4 2− ions increases the ionic strength of the solution, while Henry’s constant increases based on Eqs. (9) and (10). Therefore, the liquid-side mass transfer coefficient decreases according to Eq. (11) (Cullinane and Rochelle, 2004, 2006).
kg =
DCO2, L (k2 Cam ) HCO2, L
(11)
In addition, wet (about 45–51% moisture) and dry resins (little moisture) are used to treat spent absorbent. And the viscosities are also analyzed in Fig. 5 which shows that the viscosity is significantly reduced after the ion exchange treatment using wet and dry resins. As an example, the viscosity decreases from about 15 mPa s to 1.5 mPa s at 313 K after wet resin treatment. The diffusion coefficient, DCO2, L , then increases as the viscosity decreases according to the Eq. (7), indicating that a possible increase of the liquid-side mass transfer coefficient. 3.2.2. pH The pH of the lean liquid solution at the bottom of the stripper is shown in Fig. 6. The changes of the pH after running 100 h are small for the test without SO2 . However, the test with 200 ppm SO2 has a large pH decrease after operating 100 h because of the amine degradation and combination of SO2 with amine. The effect of pH on the overall CO2 reaction rate can be analyzed using (Sun et al., 2005): rov = kam [AM][CO2 ] + kOH− [OH− ][CO2 ]
(12)
It is obviously that the overall CO2 reaction rate is proportional to the OH− ions concentration [OH− ] and free amine concentration
3
CR1 NH3 + + CR2 NH3 + = CHCO−
(17)
3
where R1 NH2 and R2 NH2 denote the tertiary amine and additive, respectively. C1 and C2 are the amine and additive concentration, respectively. ˛CO2 is the CO2 loading. Eq. (18) instead of Eq. (17) is applied in the presence of SO2 . CR1 NH3 + + CR2 NH3 + = CHCO− + CSO2− + CSO2− 3
3
(18)
4
The three main dissociation reactions are taken into consideration and the corresponding thermodynamic equilibrium constant can be expressed as follows: K1 =
K2 =
K3 =
CR1 NH2 CH + CR
+ 1 NH 3
CR2 NH2 CH+ CR2 NH + 3
CHCO− CH+ 3
CCO2
(19)
(20)
(21)
The loss of free amine at the bottom of the stripper, including the AD, SCA and CCA, is shown in Fig. 7. The SCA and CCA are divided by a simple model mentioned above to make calculation convenient. The percent of free amine loss is defined as the lost amine concentration divided by the fresh amine concentration. By comparing two campaigns in Fig. 7, it can be found that the CCA is reduced by adding 200 ppm SO2 , particularly running 200 h later. This is because the SCA is increased gradually, and pH decreases, which is benefit to strip CO2 . Meanwhile, the AD of the campaign 3 with 200 ppm SO2 is larger than that of campaign 2 without SO2 . However, the SCA is the largest among the three parts after operating 200 h. Therefore, the SO2 is the major factor resulting in the loss of free amine. The decrease in the free amine concentration then directly influences the chemical reaction rate as indicated by
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
24
20
5.0
CO2 combines with amine for 200 ppm SO2 SO2 combines with amine for 200 ppm SO2
4.5
Degraded amine for 200 ppm SO2
4.0
Heat Stable Salts (wt%)
Percent of amine (%)
CO2 combines with amine for 0 ppm SO2 16
43
Degraded amine for 0 ppm SO2
12
8
0 ppm SO2 200 ppm SO2 After treatment for 200 ppm SO2
3.5 3.0 2.5 2.0 1.5 1.0
4
0.5
0 50
100
150
200
250
300
350
400
450
500
0.0
550
100
200
Time (h)
3.3. Ion exchange results. The ion exchange resin after percolation with an excess of NaOH solution will supply abundant OH− to the samples which will react with the bound amine as the following correlation: ¯ + H2 O ¯ → RNH2 + BA RNH3 + A− + BOH
(22)
where A− symbolizes impurity anions, such as SO3 2− , SO4 2− , HCOO− , C2 O4 2− and C2 H3 O3 − . B is the resin functional group. The impurity anions are replaced by OH− in the ion exchange process. The HSS removal efficiencies for five samples (100 h, 200 h, 300 h, 400 h and 430 h) are shown in Fig. 8. The HSS removal efficiency in Fig. 8 can be used to estimate the liquid flow rate of the original solution and design the ion exchange column (Pedruzzi et al., 2008). The average removal efficiency of formate is only 43%, but it is around 80% for glycolate, oxalate, sulfate and sulfite. The total weight percent of HSS is an important parameter influencing the CO2 capture process. The total HSS weight percents in each sample
are compared in Fig. 9. The test without SO2 has limited HSS which is only 0.14 wt% after running 500 h, which is far below the limit of 1 wt%. The limit is easily exceeded after only operating 100 h later for the sample with 200 ppm SO2 . Moreover, the total weight percent of HSS at the end of this test is about 4.5 times the limit. After treatment by the ion exchange resin, the total weight percents of HSS of all five samples with 200 ppm SO2 are lower than 1 wt%, but still higher than that of the samples without SO2 . The dynamic characteristics of the HSS removal are shown in Fig. 10. The removal efficiency generally increases slowly with the reaction time for glycolate, oxalate, sulfate and sulfite. For formate, however, the removal efficiency is almost constant during the experiment. The dynamic characteristics of the ion exchange processes for the 430 h SO2 sample are shown in Fig. 11. The pH increases with ion exchange time while the HSS weight percent decreases with reaction time. There is a dramatic change in the pH and the HSS concentration between 20 min and 30 min. After that, the pH is almost constant, while the HSS weight percent slowly decreases. Therefore, the ion exchange chemical reaction is almost
100
100h 200h 300h 400h 430h
10min 20min 30min 40min 50min 60min
90 80
Removal Efficiency (%)
Removal Efficiency (%)
500
Fig. 9. Comparison of HSS content of different conditions.
Eq. (12). It is possible that the SCA is liberated by ion exchange to recover the spent absorbents.
80
400
Time (h)
Fig. 7. Loss of free amine.
100
300
60
40
70 60 50 40 30 20
20
10 0
0
Glcolate
Formate
Oxalate
HSS Fig. 8. HSS removal efficiency.
Sulfate
Sulfite
Glcolate
Formate
Oxalate
Sulfate
Sulfite
HSS Fig. 10. Changes in HSS removal efficiencies with reaction time.
44
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
the overall gas transfer coefficient. KG of the upper line (after the resin treatment) is 1.7 times that of the lower line (before treatment). Therefore, the ion exchange can significantly improve the CO2 absorption rate of the amine aqueous solution. The same result could be obtained using (Chen et al., 2011):
1.6
10.20 10.15
1.4
10.05
1.2
pH
10.00 1.0
9.95 9.90
0.8
Heat Stable Salts (wt%)
10.10
9.85 9.80 20
30
40
50
60
The result in Figs. 5, 6, 9 and 11 show that the liquid-side mass transfer coefficient increases after the resin treatment while the gas-side mass transfer coefficient retains constant for these conditions. Accordingly, the overall mass transfer coefficient should increase according to Eq. (25). 4. Discussion
Time (min) Fig. 11. Dynamic characteristics of the ion exchange processes for the 430 h SO2 samples.
0.08
Before treatment After treatment
0.07
0.06
2
(25)
0.6 10
NCO2 (kmol/(m .h))
1 1 1 = + KG kg kg
slope=0.00215 0.05
0.04
0.03
The CO2 removal efficiency of the campaign with SO2 was about 80% after running 200 h (Gao et al., 2011c). After that, it decreased dramatically. In this work, the SO2 loading of the 200 h sample is about 0.1 mol/mol amine. At this time, SO2 combines with less free amine than CO2 and the total HSS concentration is less than 2 wt%. The AD is also changed small. After the resin treatment, the HSS concentration is less than 0.5 wt% that is half of the limit for practical use. Consequently, the appropriate operating time for reclaiming amine from the amine system is about 200 h for still keeping a high CO2 removal efficiency. Also all results show that the FGD with 90% SO2 removal efficiency is enough for the new solvent, and the rest will be removed by ion exchange. An economics analysis to adjust the suitable SO2 removal efficiency based on the FGD requirement and the amine reclamation cost will be carried out in the next step. 5. Conclusions
slope=0.00128 0.02
0.01 10
15
20
25
30
35
PCO2,b (kPa) Fig. 12. Mass transfer performance before and after the ion exchange treatment.
in equilibrium at 30 min. The repeat experiments are conducted and present good repeatability. For deeply proved an effect of ion exchange, kinetic experiments for the 430 h SO2 samples before and after resin treatment are conducted at 313 K using a wetted wall column (Liu et al., 2009). The CO2 flux is correlated with the overall gas transfer coefficient as (Dang and Rochelle, 2008)
∗ NCO2 = KG PCO2 ,b − PCO
(23)
2
where NCO2 is the CO2 flux with units of kmol/(m2 h). The CO2 equilibrium partial pressure at the gas-liquid interface, PC∗ O , will 2 approach zero since the solution loading is low. PCO2, b is the log mean CO2 partial pressure at the bulk gas side in the wetted wall column. PCO2 ,b =
PCO2 ,in − PCO2 ,out
Ln PCO2 ,in /P CO2 ,out
(24)
Thus, the overall mass transfer coefficient, KG , can be calculated as the slope of the flux in Fig. 12 versus PCO2, b . The experiments are carried out at 313 K with the results given in Fig. 12. The CO2 flux is proportional to the CO2 average partial pressure with a higher flux after the resin treatment. The slopes of the two lines represent
A pilot plant with amine-based chemical absorption process was continuously operated for about 200 h and 500 h with a feed gas of 12 vol% CO2 , 18 vol% O2 and N2 as the balance, using MEA and new solvent, respectively. In addition, 200 ppm SO2 was continuously added into new solvent and operated for 430 h. The total energy requirement of the three campaigns was analyzed. Aqueous amine solution samples were taken from the stripper bottom of the pilot plant every 100 h. The main characteristics of these samples, such as viscosity, pH value and free amine loss, were analyzed. Five solvent samples from the SO2 test were treated with an industrial grade strong base anion exchange resin from Rohm and Haas to improve the characteristics. The experimental results show that the average energy consumption of new solvent (2.96 GJ/tCO2 ) was lower than that of MEA (4.09 GJ/tCO2 ) because of lower heat of reaction and latent heat of vaporization. An addition of SO2 increased it to 3.51 GJ/tCO2 as the CO2 removal efficiency was decreased. Furthermore, the pH was reduced, the viscosity increased and more free amine was lost with more and more SO2 in the flue gas, which reduced the overall CO2 reaction rate and the overall mass transfer coefficient. The addition of SO2 increased the SO2 combines with amine and amine degradation, but decreased the CO2 combines with amine. The SO2 combines with amine became main factor after running 200 h. The resin through several runs was used to decrease the SO2 combines with amine and improve the solvent characteristics. The heat stable salts (HSS), including formate, oxalate, glycolate, sulfite and sulfate, could be reduced to less than 1 wt% by the ion exchange. The solvent viscosity after treatment was only one tenth of the original solvent at 313 K. The overall mass transfer coefficient KG after treatment was about 1.7 times that without treatment. Therefore, the flue gas desulfurization (FGD) with 90% SO2 removal efficiency combines
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
with a CO2 capture system with ion exchange equipment may be used to remove CO2 and SO2 simultaneously and economically. Acknowledgments The authors would like to acknowledge the support from the National Natural Science Foundation of China (No. 51306194), The National Natural Science Fund for Distinguished Young Scholars (No. 21425625) and the Strategic Priority Research Program of the Chinese Academy of Sciences (XDA07040403). References Ahn, H., Luberti, M., Liu, Z., Brandani, S., 2013. Process configuration studies of the amine capture process for coal-fired power plants. Int. J. Greenhouse Gas Control 16, 29–40. Bouillon, P.A., Hennes, S., Mahieux, C., 2009. ECO2 : post-combustion or oxyfuel – a comparison between coal power plants with integrated CO2 capture. Energy Procedia 1, 4015–4022. Chen, W., Liu J., Pei X., Zhang, J., Wang, X., 2008. Method for removing heat-stable salt, involves filling ion exchange resin in tower including amine sorption agent including heat-stable salt, supplying amine sorption agent for removing heat-stable salt, and regenerating resin. C.N. Patent 101, 104,1 34. Chen, X., Closmann, F., Rochelle, G.T., 2011. Accurate screening of amines by the wetted wall Column. Energy Procedia 4, 101–108. Cifre, P.G., Brechtel, K., Hoch, S., García, H., Asprion, N., Hasse, H., Scheffknecht, G., 2009. Integration of a chemical process model in a power plant modelling tool for the simulation of an amine based CO2 scrubber. Fuel 88, 2481–2488. Cullinane, J.T., Rochelle, G.T., 2004. Carbon dioxide absorption with aqueous potassium carbonate promoted by piperazine. Chem. Eng. Sci. 59 (17), 3619–3630. Cullinane, J.T., Rochelle, G.T., 2006. Kinetics of carbon dioxide absorption into aqueous potassium carbonate and piperazine. Ind. Eng. Chem. Res. 45 (8), 2531–2545. Dang, H.Y., Rochelle, G.T., 2008. CO2 absorption rate and solubility in monoethanolamine/piperazine/water. Sep. Sci. Technol. 38 (2), 337–357. Dave, N., Do, T., Puxty, G., Rowland, R., Feron, P.H.M., Attalla, M.I., 2009. CO2 capture by aqueous amines and aqueous ammonia. Energy Procedia 1, 949–954. Fang, M.X., Wang, Z., Yan, S.P., Cen, Q.G., Luo, Z.Y., 2012. CO2 desorption from rich alkanolamine solution by using membrane vacuum regeneration technology. Int. J. Greenhouse Gas Control 9, 507–521. Gao, J.B., Wang, S.J., Zhou, S., Zhao, B., Chen, C.H., 2011a. Corrosion and degradation performance of novel absorbent for CO2 capture in pilot-scale. Energy Procedia 4, 1534–1541. Gregory, J.B., Ralph, H.W., 1994. Physical solubility of carbon dioxide in aqueous alkanolamines via nitrous oxide analogy. J. Chem. Eng. Data 39 (4), 817–822. Gao, J.B., Wang, S.J., Zhou, S., Zhao, B., Chen, C.H., Xu, X.C., 2011b. Experimental study on the influence of SO2 to the CO2 capture process. Proc. CSEE 31 (5), 52–57. Gao, H.X., Zhou, L.P., Liang, Z.W., Idem, R.O., Fu, K.Y., Sema, T., Tontiwachwuthikul, P., 2014. Comparative studies of heat duty and total equivalent work of a new heat pump distillation with split flow process, conventional split flow process, and conventional baseline process for CO2 capture using monoethanolamine. Int. J. Greenhouse Gas Control 24, 87–97. Gao, J.B., Wang, S.J., Zhao, B., Qi, G.J., Chen, C.H., 2011c. Pilot-scale experimental study on the CO2 capture process with existing of SO2 : degradation, reaction rate and mass transfer. Energy Fuels 25 (12), 5802–5809. Gao, J.B., Wang, S.J., Sun, C.C., Zhao, B., Chen, C.H., 2012. Corrosion behavior of carbon steel at typical positions of an amine-based CO2 capture pilot plant. Ind. Eng. Chem. Res. 51, 6714–6721. Kumar, S., Cho, J.H., Moon, I., 2014. Ionic liquid-amine blends and CO2 BOLs: prospective solvents for natural gas sweetening and CO2 capture
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technology—a review. Int. J. Greenhouse Gas Control 20, 87–116. Knudsen, J.N., Jensen, P.J., Vilhelmsen, P.J., Biede, O., 2009. Experience with CO2 capture from coal flue gas in pilot-scale: testing of different amine solvents. Energy Procedia 1, 783–790. Keller, A.E., 1991. Acid gas removal from gas streams by alkanolamine scrubbing with removal of heat-stable salts by ion exchange. U.S. Patent 5,045,291. Liu, J.Z., Wang, S.J., Zhao, B., 2009. Absorption of carbon dioxide in aqueous ammonia. Energy Procedia 1, 933–940. Li, B., Duan, Y., Luebke, D., Morraele, B., 2013. Advances in CO2 capture technology: a patent review. Appl. Energy 102, 1439–1447. Meng, H., Zhang, S., Li, C.X., Liang, S.L., 2008. Removal of heat stable salts from aqueous solutions of N-methyldiethanolamine using a specially designed three-compartment configuration electrodialyzer. J. Membr. Sci 322, 436–440. Mathias, P.M., Reddy, S., Smith, A., Afshar, K., 2013. Thermodynamic analysis of CO2 capture solvents. Int. J. Greenhouse Gas Control 19, 262–270. Narendra, V., Anil, V., 2009. Amine system problems arising from heat stable salts and solutions to improve system performance. Fuel Process. Technol. 90, 483–489. Prashanti, B.K., Prakish, D.V., Eugeny, Y.K., 2010. Kinetics of removal of carbon dioxide by aqueous solutions of N,N-diethylethanolamine and piperazine. Environ. Sci. Technol. 44, 2138–2143. Pedruzzi, I., Borgesda, S.E.A., Rodrigues, A.E., 2008. Selection of resins, equilibrium and sorption kinetics of lactobionic acid fructose, lactose and sorbitol. Sep. Purif. Technol. 63, 600–611. Pirngruber, G.D., Guillou, F., Gomez, A., Clausse, M., 2013. A theoretical analysis of the energy consumption of post-combustion CO2 capture processes by temperature swing adsorption using solid sorbents. Int. J. Greenhouse Gas Control 14, 74–83. Rao, A.B., Ruben, E.S., 2002. A technical, economic, and environmental assessment of amine-based CO2 capture technology for power plant greenhouse gas control. Environ. Sci. Technol. 36, 4467–4475. Razi, N., Svendsen, H.F., Bolland, O., 2013. Cost and energy sensitivity analysis of absorber design in CO2 capture with MEA. Int. J. Greenhouse Gas Control 19, 331–339. Rochelle, G.T., Goff, G.S., Cullinane, J.T., Freguia, S., 2002. Research results for CO2 capture from flue gas by aqueous absorption/stripping. In: The Laurance Reid Gas Conditioning Conference, Norman, USA, pp. 131–151. Sun, W.C., Yong, C.B., Li, H.M., 2005. Kinetics of the absorption of carbon dioxide into mixed aqueous solutions of 2-amino-2-methyl-1-propanol and piperazine. Chem. Eng. Sci. 60, 503–516. Versteeg, G.F., van Swaaij, W.P.M., 1988. Solubility and diffusivity of acid gases (CO2 , N2 O) in aqueous alkanolamine solutions. J. Chem. Eng. Data 33, 29–34. Warudkar, S.S., Cox, K.R., Wong, M.S., Hirasaki, G.J., 2013a. Influence of stripper operating parameters on the performance of amine absorption systems for post-combustion carbon capture: part I. High pressure strippers. Int. J. Greenhouse Gas Control 16, 342–350. Warudkar, S.S., Cox, K.R., Wong, M.S., Hirasaki, G.J., 2013b. Influence of stripper operating parameters on the performance of amine absorption systems for post-combustion carbon capture: part II. Vacuum strippers. Int. J. Greenhouse Gas Control 16, 351–360. Xu, F., Gao, H.S., Dong, H.F., Wang, Z.L., Zhang, X.P., Ren, B.Z., Zhang, S.J., 2014. Solubility of CO2 in aqueous mixtures of monoethanolamine and dicyanamide-based ionic liquids. Fluid Phase Equilibria 365, 80–87. Xu, Q., Rochelle, G.T., 2009. Solvent reclaiming by crystallization of potassium sulfate. Energy Procedia 1, 1205–1212. Yang, J., Yu, X.H., Yan, J.Y., Tu, S.-T., Dahlquist, E., 2013. Effects of SO2 on CO2 capture using a hollow fiber membrane contactor. Appl. Energy 112, 755–764. Yu, H., Morgan, S., Allport, A., Cottrell, A., Do, T., McGregor, J., Wardhaugh, L., Feron, P., 2011. Results from trialling aqueous NH3 based post-combustion capture in a pilot plant at Munmorah power station: absorption. Chem. Eng. Res. Des. 89, 1204–1215. Zhang, X.P., Zhang, X.C., Dong, H.F., Zhao, Z.J., Zhang, S.J., Huang, Y., 2012. Carbon capture with ionic liquids: overview and progress. Energy Environ. Sci. 5, 6668–6681.
Contents lists available at ScienceDirect
International Journal of Greenhouse Gas Control journal homepage: www.elsevier.com/locate/ijggc
Effect of SO2 on the amine-based CO2 capture solvent and improvement using ion exchange resins Jubao Gao a,∗ , Shujuan Wang b,∗∗ , Jian Wang a,c , Lingdi Cao a,c , Shiwei Tang a , Yu Xia a a Beijing Key Laboratory of Ionic Liquids Clean Process, Key Laboratory of Green Process and Engineering, State Key Laboratory of Multiphase Complex Systems, Institute of Process Engineering, Chinese Academy of Sciences, Beijing 100190, China b Key Laboratory for Thermal Science and Power Engineering of Ministry of Education Department of Thermal Engineering, Tsinghua University, Beijing 100084, China c School of Chemistry and Chemical engineering, University of Chinese Academy of Sciences, Beijing 100049, China
a r t i c l e
i n f o
Article history: Received 28 July 2014 Received in revised form 12 January 2015 Accepted 4 March 2015 Keywords: Carbon dioxide Sulfur dioxide Pilot plant Heat stable salts Ion exchange resin Energy consumption
a b s t r a c t An amine-based CO2 capture solvent is developed to reduce energy requirement. Three campaigns, with 12 vol% CO2 and 18 vol% O2 , are conducted in a CO2 capture pilot plant to evaluate energy consumption and influence of SO2 on solvent characteristics, under the condition of prolonged operation. Monoethanolamine (MEA) is used in one campaign, and the other two campaigns employed new solvent, with and without 200 ppm SO2 in the simulated flue gas. The lost of free amine, including CO2 combines with amine (CCA), SO2 combines with amine (SCA) and amine degradation (AD), is analyzed. Experimental results show that the average energy consumption of MEA and novel absorbent is about 4.09 GJ/tCO2 and 2.96 GJ/tCO2 , respectively. This value climbs to 3.51 GJ/tCO2 for novel absorbent, with 200 ppm SO2 . Meanwhile, the pH is reduced, the viscosity increases and free amine is lost more, which reduces the overall CO2 reaction rate and the overall mass transfer coefficient. The presence of SO2 increases SCA and AD, and decreases CCA. The SCA becomes main factor after running 200 h. An industrial strong-alkaline anion exchange resin through several runs is used to decrease the SCA and improve the solvent characteristics. The heat stable salts (HSS), including formate, oxalate, glycolate, sulfite and sulfate, can be reduced to less than 1 wt% by using the ion exchange resin. The solvent viscosity after treatment is only one tenth of the original solvent at 313 K. The overall mass transfer coefficient KG after treatment is about 1.7 times that without treatment. © 2015 Elsevier Ltd. All rights reserved.
1. Introduction Amine scrubbing systems have been used to separate CO2 from natural gas in the chemical industry for several decades. The flue gas from coal-fired power plants, however, is quite different from that of natural gas, with large flow rates, high temperatures, low CO2 partial pressures, and the existence of SO2 , NOx and fly ash in the flow stream (Xu et al., 2014; Zhang et al., 2012). Therefore, removing CO2 from power plants results in significant energy consumption and reduction of electricity output (Bouillon et al., 2009; Knudsen et al., 2009; Razi et al., 2013). Many ways have been reported to minimize the CO2 capture energy requirement (Ahn et al., 2013; Fang et al., 2012; Gao et al., 2014; Warudkar et al., 2013a,b), solvent screening is another efficient method to improve
∗ Corresponding author. Tel.: +86 10 82544875; fax: +86 10 82544875. ∗∗ Corresponding author. Tel.: +86 10 62788668; fax: +86 10 62770209. E-mail addresses: [email protected] (J. Gao), [email protected] (S. Wang). http://dx.doi.org/10.1016/j.ijggc.2015.03.001 1750-5836/© 2015 Elsevier Ltd. All rights reserved.
the energy efficiency (Gao et al., 2011a; Kumar et al., 2014; Mathias et al., 2013; Pirngruber et al., 2013; Yu et al., 2011; Zhang et al., 2012). Recently, Li et al. (2013) found that about 380 patents have been published on CO2 capture by solvents. In this work, a novel blended amine solvent was developed and assessed in a pilot plant, with the long-term running. In addition, SO2 in the flue gas is one of the main factors causing free amine lost, which may lead to rise of CO2 capture energy consumption. In previous work, experiments have proved that SO2 in the flue gas results in serious amine degradation (AD), with many heat stable salts (HSS) produced (Gao et al., 2011a,b,c; Yang et al., 2013). Then, foaming and corrosion/erosion can be caused by the high levels of HSS (>1 wt% as amine) (Gao et al., 2012; Narendra and Anil, 2009). However, it is rarely to analysis the effect of SO2 on the CO2 combines with amine (CCA) and SO2 combines with amine (SCA), which leads to free amine lost. The upgraded FGD with more than 99% SO2 removal efficiency before CO2 capture was suggested by Rao and Ruben (2002) to solve the problems. But HSS and impurities are still accumulated in amine solutions. Hence, amine
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
Nomenclature Cam DCO2, L
amine concentration in the liquid phase (kmol/m3 ) CO2 diffusion coefficient in the amine solution
DCO2, w HCO2, w
(m2 /s) CO2 diffusion coefficient in water (m2 /s) Henry’s law coefficient of CO2 in the amine solution
Qstrip Qsens Qdes Qloss QT rov
(kPa m3 /kmol) second order rate constant of amine (m3 /kmol s) overall mass-transfer coefficient (kmol/(m2 h kPa)) gas-side mass-transfer coefficient (kmol/(m2 h kPa)) liquid-side mass-transfer coefficient (kmol/(m2 h kPa)) mass-transfer flux of the absorbed component CO2 (kmol/(m2 h)) average CO2 partial pressure (kPa) equilibrium CO2 partial pressure (kPa) CO2 partial pressure at the wetted wall column inlet (kPa) CO2 partial pressure at the wetted wall column outlet (kPa) latent heat of vaporization (kJ/mol CO2 ) sensible heat (kJ/mol CO2 ) heat of reaction (kJ/mol CO2 ) heat of loss (GJ/tCO2 ) total energy consumption (GJ/tCO2 ) overall CO2 reaction rate (kmol/m3 s)
Greeks L W ˛CO2
amine solution viscosity (mPa s) water viscosity (mPa s) CO2 loading (mol/mol)
k2 KG kg k g NCO2 PCO2, b P ∗ CO2 PCO2, in PCO2, out
reclamation is essential during CO2 capture to control the levels of HSS and remove impurities. Amine reclamation can be used to prevent potential secondary pollution with HSS removal and increases in the free amine concentration in a continuous absorption/regeneration process. Three prime technologies used commercially for solvent reclamation as thermal reclamation, electrodialysis, and ion-exchange. Thermal reclamation at the bottom of the stripper (Fig. 1) can result in higher reclamation costs and potential degradation of the amine due to the high temperatures (Xu and Rochelle, 2009). This is a useful method to remove non-volatile species. Electrodialysis is used to remove anions and cations, however, it can cause absorbent losses. Moreover, the ion-exchange membrane is relatively expensive especially in the presence of fly ash (Meng et al., 2008). Ion-exchange systems acts like electrodialysis disable to remove nonionized impurities. And ion-exchange resins must be used at low temperatures to reduce resin lost. Hence, ion-exchange systems all normally installed at the absorber inlet (Fig. 1). Ion-exchange has been widely used in coal-fired power plants to treat boiler feed-water, thus, it may be more easily accepted in coal-fired power plants for CO2 capture. The ion-exchange technology for removing HSS was patented by Keller (1991) and Chen et al. (2008). However, there is little information on the HSS type and removal efficiency. Meanwhile, these patents were only applied in removing H2 S and CO2 from natural gas. Although there is substantial information on novel solvents and methods of saving energy requirement, the information on the energy consumption analysis from pilot plant and effects of SO2 on absorbents, in particulate, improved methods are scant. Three
39
Table 1 Anion exchange resin properties.a Total exchange capacity (eq/L)
Moisture holding capacity (%)
Shipping weight (g/L)
Particle mean size (mm)
Uniformity coefficient
1.25
45–51
680
0.664
1.6
a
Data obtained from supplier except for particle size.
campaigns have been conducted in a pilot plant using simulated flue gas to assess CO2 capture energy consumption and effect of SO2 on CO2 capture solvent in this study. The main characteristics of amine aqueous solutions, such as viscosity, pH and bound amine, were analyzed after long-term operation, with and without 200 ppm SO2 . The AD, CCA and SCA of two campaigns were compared. An industrial grade anion exchange resin was selected to improve the characteristics of the amine-based CO2 absorbent. 2. Experimental 2.1. Materials The chemical reagents are industry grade, with concentrations of more than 98%. CO2 , SO2 , O2 and N2 , with more than 99.9% concentration, are used to simulate coal-fired flue gas. New solvent is a blend of a tertiary amine and additives. Aqueous amine solution samples are taken from the stripper bottom of the pilot plant every 100 h (Gao et al., 2011a,b,c). The character of four types of strong-alkaline and two types of weak-alkaline anion-exchange resin is compared, and the best is selected. Five solvent samples from the SO2 test are treated by the selected industrial grade strong base anion exchange resin from Rohm and Haas. The resin is made of pale yellow, translucent spherical beads with the main component being styrene divinylbenzene copolymer matrix with a dimethyl ethanol ammonium functional group. The anion is chloride as shipped and is converted to hydroxyl by percolation with an excess of 4 M NaOH solution and washed with deionized water. Table 1 shows the typical properties of the anion exchange resin. 2.2. Pilot plant operating conditions Three campaigns run in a CO2 capture pilot plant. Some important parameters are shown in Table 2. Two baseline campaigns operated, respectively, for 200 h and 500 h using 30 wt% MEA and novel solvent, with a feed gas of 12 vol% CO2 , 18 vol% O2 and N2 as the balance, and another campaign using the novel solvent, feed with an additional 200 ppm SO2 , operated for 430 h. The pilot plant is same as conventional amine-based CO2 recovery plant except for recycled flue gas from stripper top to absorber inlet (Gao et al., 2011a,c). The average flue gas flow rate of the two campaigns is about 86 Nm3 /h, and the average liquid solvent flow rate is around 0.6 m3 /h. The heat input to the stripper is supplied by electrical heater with the power of 60 kW. No fresh solvent and water are added into the system during the two campaigns. Thus, the CO2 removal efficiency is changed during the two campaigns because of free amine lost. The average temperature at the bottom of absorber and stripper is respectively around 46 ◦ C and 116 ◦ C. More details can be found from our previous work (Gao et al., 2011c). 2.3. Ion exchange procedure The resin is first percolated and held for 24 h with an excess of 4 M NaOH solution. Then it is washed three times and held for 3 h in deionized water. The pretreated resin is then put into ten 1 L beakers with 170 ml in each. 100 ml solvent samples are fed into
40
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
Fig. 1. A general flow diagram of CO2 capture with flue gas pretreatment and amine reclamation. Table 2 Experimental conditions. Items Unit
CO2 Vol%
O2 Vol%
SO2 ×10−6
Amine concentration mol/L
Flue gas flow rate Nm3 /h
Operating time h
Campaign 1 Campaign 2 Campaign 3
12.89 11.89 11.86
17.4 18.2 17.8
0 0 200
4.90 4.64 4.57
85.35 86.12 82.83
200 500 430
2.4. Analysis methods Dionex ICS-1000 and DX120 ion chromatographics are used to analyze the anions and cations. CO2 loading is determined by a total organic carbon analyzer (TOC-5000, Shimadzu). The solution viscosity is measured by a viscometer (SV-10, AND). The pH is measured on an automatic potentiometric titrator (809 Titrando, Metrohm, Switzerland). A wetted wall column with an interfacial area of about 41.45 cm2 is used to measure the kinetics of the 430 h samples before and after treatment. The column is described by Liu et al., (2009).
6
New Solvent MEA energy consumption ( GJ/t CO2)
each beaker. Five different samples from the SO2 tests (100 h, 200 h, 300 h, 400 h and 430 h samples) are held for 60 minutes in each beaker. The 430 h samples are fed into five beakers, each one lasting 10 min longer than the previous one (10, 20, 30, 40 and 50 min after reaction at room temperature). The mixture is stirred at a constant rate before the 10 ml samples is taken for analysis.
5
4
3
2 0
100
200
300
400
500
Time (h) Fig. 2. Energy consumption of different solvent.
3. Results and discussion 3.1. Energy consumption The total energy consumption, including energy loss of the whole CO2 capture system, the sensible heat, the latent heat of vaporization and the overall enthalpies of absorption, is determined experimentally by the following correlation: QT =
(tin − tout ) Coil oil Voil qCO2 ∗ 1000
(1)
where, tin and tout denote temperature of oil entering and leaving the reboiler, respectively. Coil , oil and Voil are the heat capacity, the density and the volume flow rate of oil with unit of kJ/(kg K), kg/m3 and m3 /h, respectively. qCO2 reprensents the mass flow rate of gas CO2 at the stripper top with unit of kg/h.
Results of the total energy consumption are shown in Fig. 2 and Fig. 3. It can be seen from Fig 2 that the total average energy consumption for 30 wt% MEA is about 4.09 GJ/tCO2 (3.53–4.88 GJ/tCO2 ), which is in good agreement with simulated results of Dave et al. (2009) and Cifre et al. (2009), but about 10% higher than that reported by Knudsen et al. (2009). For new solvent, however, it is about 2.96 GJ/tCO2 (2.46–3.15 GJ/tCO2 ), which is about 27% lower than that of MEA. After adding SO2 into the CO2 capture system (Fig. 3), this value has been increased to 3.51 GJ/tCO2 (1.69–5.05 GJ/tCO2 ). Before running of 150 h, campaign 3 has a comparable value to campaign 2. After running of 150 h, nevertheless, campaign 3 is gradually higher than campaign 2 due to the reduction of the CO2 removal efficiency (Gao et al., 2011c).
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
140
7.0 6.5
0ppm SO2
130
6.0
200ppm SO2
120
sensible heat heat of reaction heat of vaporization loss energy
110
5.5
100
5.0
Energy distribution %
energy consumption ( GJ/t CO2)
41
4.5 4.0 3.5 3.0 2.5
90 80 70 60 50 40
2.0
30
1.5
20 10
1.0 0
50
100
150
200
250
300
350
400
450
500
0
MEA
Time (h)
New Solvent
Absorption Solvent Fig. 3. Effect of SO2 on energy consumption. Fig. 4. Distribution of the total energy.
In addition, each component (loss heat, latent heat of vaporization, sensible heat and reaction enthalpy) can be expanded separately from the pilot plant experiments as following (Rochelle et al., 2002): QT = Qloss + Qsens + Qdes + Qstrip
(2)
Qdes = −Habs
(3)
Qsens
Qstrip =
(4)
2
Trich,str , ˛rich
n
Qloss = 3.6
(5)
18
Fresh solvent
14
8
0ppm SO2
6
[˛ti Ai (tsi − ta )] /qCO2 /1000
(6)
3.2. Influence of SO2 on the absorbent The HSS contents of the five SO2 samples are shown in Table 3. Most of the HSS is sulfates and sulfites because SO2 is absorbed conTable 3 HSS from CO2 capture pilot plant. Glycolate
Formate
Oxalate
105.7 170.2 271.2 396.8 375.4
22.1 78.4 157.9 248.4 335.5
Sulfate
Sulfite
mg/L 322.7 680.6 965.5 1302.2 1706.9
4 Treatment with 2 dry resins
Treatment with wet resins
0
where ˛ti and Ai are the total heat transfer coefficient and surface area of the i-th equipment with unit of w/(m2 K) and m2 , respectively. tsi and ta denote the surface average temperature of the i-th equipment and atmosphere temperature, respectively. The distributions of the four energy components for campaign 1 and 2 are illustrated in Fig. 4. It can be found that MEA has a comparable loss energy and sensible heat to new solvent. However, the heat of reaction and the latent heat of vaporization of new solvent are all lower than those of MEA, which implies that regenerating the new solvent requires low reboiler heat duty to break the chemical bounds between CO2 and amine and strip CO2 .
100 h 200 h 300 h 400 h 430 h
20
10
HHvap 2O
i=1
Sample
22
12
sat PH Trich,str xH2 O O ∗ PCO 2
200ppm SO2
24
16
CP = (˛rich −˛lean ) Cam
26
5007.5 12,144.0 19,454.7 25,723.2 33,837.0
5145.4 5118.2 5332.1 6095.5 8650.4
300
305
310
313K
315
320
325
330
335
340
Temperature (K) Fig. 5. Solvent viscosities.
tinuously. The effect of SO2 on the solvents is shown in the following sections. 3.2.1. Viscosity The viscosities of fresh amine, the 500 h sample and the 430 h sample are measured. As shown in Fig. 5, the solvent viscosities of the 500 h samples (0 ppm SO2 ) are similar to that of the fresh amine. The highest viscosity of the sample (200 ppm SO2 ), however, is observed. In general, amine concentration is proportional to the viscosity. Thus, two spent absorbents should cut down on the viscosity because of amine degradation, which indicates that the formation of HSS increases the viscosity especially in the presence of SO2 . Furthermore, the CO2 diffusion coefficient in solution DCO2, L can be calculated from the viscosity as (Prashanti et al., 2010; Versteeg and van Swaaij, 1988): DCO2, L = DCO2, W
0.8 W
(7)
L
DCO2, W = 2.35 × 10−6 exp(
−2119 ) T
(8)
42
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
[AM]. The former can be neglected because of rather low concentration. But the reduction of [OH− ] indicates the decrease of [AM], which results in the overall CO2 reaction rate decreases. The pH of the five samples before and after treatment is shown in Fig. 6. The pH increases significantly after the ion exchange resin treatment, which liberates part of bound amine and leads to an increase in the overall CO2 reaction rate as indicated by Eq. (12). At the same time, most of the impurity anions are removed by the resin so that the ionic strength of the solution decreases. As a result, Henry’s constant for the electrolyte solution is also reduced. All these results show an increase in the liquid-side mass transfer after the ion exchange treatment.
12.0
Before treatment for CO2 +200ppm SO2 CO2 only
11.5
After treatment for CO2 +200ppm SO2
11.0
pH
10.5
10.0
9.5
3.2.3. Free amine As an acidic gas, SO2 will combine with free amine, which will reduce the free amine concentration:
9.0 0
100
200
300
400
500
Time (h)
The Henry’s law constant of CO2 is obtained using (Gregory and Ralph, 1994):
I = (0.5)
Ci Zi2
(13) HCOO− ,
Fig. 6. pH at the bottom of the stripper.
log HCO2, L = h+ + h− + hg I + log (H∗ )
SO2 + H2 O + 2RNH2 → 2RNH3 + + SO3 2−
(9) (10)
2−
In effect, the organic anions, such as C2 O4 and C2 H3 O3 − , are also combined with free amine. Their influence is not considered in this study because of low concentration. The effect of H+ and OH− on the mass balance and charge balance are ignored because of rather low concentration. The forms of the absorbed CO2 are regarded as HCO3 − since the CO2 and CO3 2− are also small. The mass balance and charge balance equations can be approximated as follows: CR1 NH2 + CR
+ 1 NH3
= C1
(14)
where HCO2, L is Henry’s constant for the electrolyte solution and
CR2 NH2 + CR2 NH3 + = C2
(15)
H* is for the pure molecular solvent. h+ and h− are the van Krevelen coefficients for the cations and anions in the solution and is the coefficient for the dissolved gas. I is the ionic strength of the solution. Ci is the concentration of species i and Zi is the charge. According to Eqs. (7) and (8) DCO2, L decreases as the viscosity
CHCO− = (C1 + C2 )˛CO2
(16)
increases. And the accumulation of SO3 2− and SO4 2− ions increases the ionic strength of the solution, while Henry’s constant increases based on Eqs. (9) and (10). Therefore, the liquid-side mass transfer coefficient decreases according to Eq. (11) (Cullinane and Rochelle, 2004, 2006).
kg =
DCO2, L (k2 Cam ) HCO2, L
(11)
In addition, wet (about 45–51% moisture) and dry resins (little moisture) are used to treat spent absorbent. And the viscosities are also analyzed in Fig. 5 which shows that the viscosity is significantly reduced after the ion exchange treatment using wet and dry resins. As an example, the viscosity decreases from about 15 mPa s to 1.5 mPa s at 313 K after wet resin treatment. The diffusion coefficient, DCO2, L , then increases as the viscosity decreases according to the Eq. (7), indicating that a possible increase of the liquid-side mass transfer coefficient. 3.2.2. pH The pH of the lean liquid solution at the bottom of the stripper is shown in Fig. 6. The changes of the pH after running 100 h are small for the test without SO2 . However, the test with 200 ppm SO2 has a large pH decrease after operating 100 h because of the amine degradation and combination of SO2 with amine. The effect of pH on the overall CO2 reaction rate can be analyzed using (Sun et al., 2005): rov = kam [AM][CO2 ] + kOH− [OH− ][CO2 ]
(12)
It is obviously that the overall CO2 reaction rate is proportional to the OH− ions concentration [OH− ] and free amine concentration
3
CR1 NH3 + + CR2 NH3 + = CHCO−
(17)
3
where R1 NH2 and R2 NH2 denote the tertiary amine and additive, respectively. C1 and C2 are the amine and additive concentration, respectively. ˛CO2 is the CO2 loading. Eq. (18) instead of Eq. (17) is applied in the presence of SO2 . CR1 NH3 + + CR2 NH3 + = CHCO− + CSO2− + CSO2− 3
3
(18)
4
The three main dissociation reactions are taken into consideration and the corresponding thermodynamic equilibrium constant can be expressed as follows: K1 =
K2 =
K3 =
CR1 NH2 CH + CR
+ 1 NH 3
CR2 NH2 CH+ CR2 NH + 3
CHCO− CH+ 3
CCO2
(19)
(20)
(21)
The loss of free amine at the bottom of the stripper, including the AD, SCA and CCA, is shown in Fig. 7. The SCA and CCA are divided by a simple model mentioned above to make calculation convenient. The percent of free amine loss is defined as the lost amine concentration divided by the fresh amine concentration. By comparing two campaigns in Fig. 7, it can be found that the CCA is reduced by adding 200 ppm SO2 , particularly running 200 h later. This is because the SCA is increased gradually, and pH decreases, which is benefit to strip CO2 . Meanwhile, the AD of the campaign 3 with 200 ppm SO2 is larger than that of campaign 2 without SO2 . However, the SCA is the largest among the three parts after operating 200 h. Therefore, the SO2 is the major factor resulting in the loss of free amine. The decrease in the free amine concentration then directly influences the chemical reaction rate as indicated by
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
24
20
5.0
CO2 combines with amine for 200 ppm SO2 SO2 combines with amine for 200 ppm SO2
4.5
Degraded amine for 200 ppm SO2
4.0
Heat Stable Salts (wt%)
Percent of amine (%)
CO2 combines with amine for 0 ppm SO2 16
43
Degraded amine for 0 ppm SO2
12
8
0 ppm SO2 200 ppm SO2 After treatment for 200 ppm SO2
3.5 3.0 2.5 2.0 1.5 1.0
4
0.5
0 50
100
150
200
250
300
350
400
450
500
0.0
550
100
200
Time (h)
3.3. Ion exchange results. The ion exchange resin after percolation with an excess of NaOH solution will supply abundant OH− to the samples which will react with the bound amine as the following correlation: ¯ + H2 O ¯ → RNH2 + BA RNH3 + A− + BOH
(22)
where A− symbolizes impurity anions, such as SO3 2− , SO4 2− , HCOO− , C2 O4 2− and C2 H3 O3 − . B is the resin functional group. The impurity anions are replaced by OH− in the ion exchange process. The HSS removal efficiencies for five samples (100 h, 200 h, 300 h, 400 h and 430 h) are shown in Fig. 8. The HSS removal efficiency in Fig. 8 can be used to estimate the liquid flow rate of the original solution and design the ion exchange column (Pedruzzi et al., 2008). The average removal efficiency of formate is only 43%, but it is around 80% for glycolate, oxalate, sulfate and sulfite. The total weight percent of HSS is an important parameter influencing the CO2 capture process. The total HSS weight percents in each sample
are compared in Fig. 9. The test without SO2 has limited HSS which is only 0.14 wt% after running 500 h, which is far below the limit of 1 wt%. The limit is easily exceeded after only operating 100 h later for the sample with 200 ppm SO2 . Moreover, the total weight percent of HSS at the end of this test is about 4.5 times the limit. After treatment by the ion exchange resin, the total weight percents of HSS of all five samples with 200 ppm SO2 are lower than 1 wt%, but still higher than that of the samples without SO2 . The dynamic characteristics of the HSS removal are shown in Fig. 10. The removal efficiency generally increases slowly with the reaction time for glycolate, oxalate, sulfate and sulfite. For formate, however, the removal efficiency is almost constant during the experiment. The dynamic characteristics of the ion exchange processes for the 430 h SO2 sample are shown in Fig. 11. The pH increases with ion exchange time while the HSS weight percent decreases with reaction time. There is a dramatic change in the pH and the HSS concentration between 20 min and 30 min. After that, the pH is almost constant, while the HSS weight percent slowly decreases. Therefore, the ion exchange chemical reaction is almost
100
100h 200h 300h 400h 430h
10min 20min 30min 40min 50min 60min
90 80
Removal Efficiency (%)
Removal Efficiency (%)
500
Fig. 9. Comparison of HSS content of different conditions.
Eq. (12). It is possible that the SCA is liberated by ion exchange to recover the spent absorbents.
80
400
Time (h)
Fig. 7. Loss of free amine.
100
300
60
40
70 60 50 40 30 20
20
10 0
0
Glcolate
Formate
Oxalate
HSS Fig. 8. HSS removal efficiency.
Sulfate
Sulfite
Glcolate
Formate
Oxalate
Sulfate
Sulfite
HSS Fig. 10. Changes in HSS removal efficiencies with reaction time.
44
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
the overall gas transfer coefficient. KG of the upper line (after the resin treatment) is 1.7 times that of the lower line (before treatment). Therefore, the ion exchange can significantly improve the CO2 absorption rate of the amine aqueous solution. The same result could be obtained using (Chen et al., 2011):
1.6
10.20 10.15
1.4
10.05
1.2
pH
10.00 1.0
9.95 9.90
0.8
Heat Stable Salts (wt%)
10.10
9.85 9.80 20
30
40
50
60
The result in Figs. 5, 6, 9 and 11 show that the liquid-side mass transfer coefficient increases after the resin treatment while the gas-side mass transfer coefficient retains constant for these conditions. Accordingly, the overall mass transfer coefficient should increase according to Eq. (25). 4. Discussion
Time (min) Fig. 11. Dynamic characteristics of the ion exchange processes for the 430 h SO2 samples.
0.08
Before treatment After treatment
0.07
0.06
2
(25)
0.6 10
NCO2 (kmol/(m .h))
1 1 1 = + KG kg kg
slope=0.00215 0.05
0.04
0.03
The CO2 removal efficiency of the campaign with SO2 was about 80% after running 200 h (Gao et al., 2011c). After that, it decreased dramatically. In this work, the SO2 loading of the 200 h sample is about 0.1 mol/mol amine. At this time, SO2 combines with less free amine than CO2 and the total HSS concentration is less than 2 wt%. The AD is also changed small. After the resin treatment, the HSS concentration is less than 0.5 wt% that is half of the limit for practical use. Consequently, the appropriate operating time for reclaiming amine from the amine system is about 200 h for still keeping a high CO2 removal efficiency. Also all results show that the FGD with 90% SO2 removal efficiency is enough for the new solvent, and the rest will be removed by ion exchange. An economics analysis to adjust the suitable SO2 removal efficiency based on the FGD requirement and the amine reclamation cost will be carried out in the next step. 5. Conclusions
slope=0.00128 0.02
0.01 10
15
20
25
30
35
PCO2,b (kPa) Fig. 12. Mass transfer performance before and after the ion exchange treatment.
in equilibrium at 30 min. The repeat experiments are conducted and present good repeatability. For deeply proved an effect of ion exchange, kinetic experiments for the 430 h SO2 samples before and after resin treatment are conducted at 313 K using a wetted wall column (Liu et al., 2009). The CO2 flux is correlated with the overall gas transfer coefficient as (Dang and Rochelle, 2008)
∗ NCO2 = KG PCO2 ,b − PCO
(23)
2
where NCO2 is the CO2 flux with units of kmol/(m2 h). The CO2 equilibrium partial pressure at the gas-liquid interface, PC∗ O , will 2 approach zero since the solution loading is low. PCO2, b is the log mean CO2 partial pressure at the bulk gas side in the wetted wall column. PCO2 ,b =
PCO2 ,in − PCO2 ,out
Ln PCO2 ,in /P CO2 ,out
(24)
Thus, the overall mass transfer coefficient, KG , can be calculated as the slope of the flux in Fig. 12 versus PCO2, b . The experiments are carried out at 313 K with the results given in Fig. 12. The CO2 flux is proportional to the CO2 average partial pressure with a higher flux after the resin treatment. The slopes of the two lines represent
A pilot plant with amine-based chemical absorption process was continuously operated for about 200 h and 500 h with a feed gas of 12 vol% CO2 , 18 vol% O2 and N2 as the balance, using MEA and new solvent, respectively. In addition, 200 ppm SO2 was continuously added into new solvent and operated for 430 h. The total energy requirement of the three campaigns was analyzed. Aqueous amine solution samples were taken from the stripper bottom of the pilot plant every 100 h. The main characteristics of these samples, such as viscosity, pH value and free amine loss, were analyzed. Five solvent samples from the SO2 test were treated with an industrial grade strong base anion exchange resin from Rohm and Haas to improve the characteristics. The experimental results show that the average energy consumption of new solvent (2.96 GJ/tCO2 ) was lower than that of MEA (4.09 GJ/tCO2 ) because of lower heat of reaction and latent heat of vaporization. An addition of SO2 increased it to 3.51 GJ/tCO2 as the CO2 removal efficiency was decreased. Furthermore, the pH was reduced, the viscosity increased and more free amine was lost with more and more SO2 in the flue gas, which reduced the overall CO2 reaction rate and the overall mass transfer coefficient. The addition of SO2 increased the SO2 combines with amine and amine degradation, but decreased the CO2 combines with amine. The SO2 combines with amine became main factor after running 200 h. The resin through several runs was used to decrease the SO2 combines with amine and improve the solvent characteristics. The heat stable salts (HSS), including formate, oxalate, glycolate, sulfite and sulfate, could be reduced to less than 1 wt% by the ion exchange. The solvent viscosity after treatment was only one tenth of the original solvent at 313 K. The overall mass transfer coefficient KG after treatment was about 1.7 times that without treatment. Therefore, the flue gas desulfurization (FGD) with 90% SO2 removal efficiency combines
J. Gao et al. / International Journal of Greenhouse Gas Control 37 (2015) 38–45
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